Catalyzed alkylation, alkylation catalysts, and methods of making alkylation catalysts

ABSTRACT

Improved alkylation catalysts, alkylation methods, and methods of making alkylation catalysts are described. The alkylation method comprises reaction over a solid acid, zeolite-based catalyst and can be conducted for relatively long periods at steady state conditions. The alkylation catalyst comprises a crystalline zeolite structure, a Si/Al molar ratio of 20 or less, less than 0.5 weight percent alkali metals, and further having a characteristic catalyst life property. Some catalysts may contain rare earth elements in the range of 10 to 35 wt %. One method of making a catalyst includes a calcination step following exchange of the rare earth element(s) conducted at a temperature of at least 575° C. to stabilize the resulting structure followed by an deammoniation treatment. An improved method of deammoniation uses low temperature oxidation.

RELATED APPLICATION

This application is a divisional of U.S. patent application Ser. No.15/190,063 filed Jun. 22, 2016, now U.S. Pat. No. 10,464,863 and claimsthe priority benefit of U.S. Provisional Patent Application Ser. No.62/183,037, filed 22 Jun. 2015.

INTRODUCTION

The term “alkylate” refers to the reaction product of isobutane witholefins. High-octane alkylate is the cleanest gasoline blending streamproduced in a refinery. Alkylate is an ideal clean fuel componentbecause pollution from alkylate is low and alkylate has low toxicity.Alkylate has been blended into gasoline for decades to improve octaneand thus the antiknock properties of gasoline. Alkylate's importance torefiners continues to grow. It makes up about 13%, or more than 12million barrels per day, of current North American fuel. For the refinerstriving to meet the stricter motor fuel specifications being mandatedon an expanding worldwide basis, investment in alkylation capacity canlead to enhanced refinery flexibility and profitability. Alkylate ismade currently using liquid acid catalysts. Refiners typically useeither hydrofluoric acid (HF), which can be deadly if spilled, orsulfuric acid (H₂SO₄), which also potentially is harmful andincreasingly costly to recycle.

In an alkylation reaction, light olefins are reacted with iso-paraffins(typically iso-butane) in the presence of a strong acid catalyst. Thealkylation of isobutane with C₂-C₅ olefins involves a series ofconsecutive and simultaneous reactions occurring through carbocationintermediates. The first step is the addition of a proton to theisobutane to form a tert-butyl, or t-butyl cation. The t-butyl cationthen is added to an olefin to give the corresponding C₈ carbocation.These C₈ carbocations may isomerize via hydride transfer and methylshifts to form more stable cations. Then the C₈ cations undergo rapidhydride transfer with isobutane, to form the desired iso-octanemolecule, and the t-butyl cation is regenerated to perpetuate the chainsequence.

Unfortunately, these are not the only reactions occurring duringalkylation. There are a number of secondary reactions that, in general,tend to reduce the quality of the alkylate. Polymerization results fromthe addition of a second olefin to the C₈+ cation formed in the primaryreaction, thereby forming cations with more than 8 carbon atoms, such asC₁₂+ cations. The C₁₂+ cation can continue to react with an olefin toform a larger cation. The successive addition of olefins tocarbocations, or olefin polymerization, is believed generally to be theprimary route to catalyst deactivation. The olefin addition reactionsometimes is referred to as a polymerization step, while thehydride-transfer reaction is denoted as the main alkylation reaction.The polymerization reaction results in the formation of “coke”. Theheavier alkylate molecules can then crack over the acid sites to formlighter C₅-C₇ hydrocarbons. As a result, alkylate consists of paraffinicmolecules from light iso-pentane (C₅H₁₂) to heavier (C₁₂H₂₆ and larger)hydrocarbons.

Solid acid catalysts have been investigated as alternatives to liquidcatalysts for nearly 30 years. Some of these catalysts include AlCl₃;platinum compounds; heteropolyacids, such as tungstates; and liquidacids immobilized on silica, polymers, or other solid supports. Naturalor artificial zeolites also have been used. Solid acid catalysts can betuned to improve selectivity and reduce production costs, but they tendto deactivate rapidly under alkylation reaction conditions through twomechanisms: 1) “Coke” formation on active sites from olefinpolymerization reaction and 2) Pore-mouth plugging by heavy alkylatemolecules The heavy hydrocarbons tend to plug the pore structure ofsolid catalysts, thereby reducing access to acidic sites.

There has been great interest in developing improved solid acidalkylation catalysts. For example, Japanese Patent Application No.1-245853, U.S. Pat. Nos. 3,962,133 and 4,116,880, and United KingdomPatent Nos. 1,432,720 and 1,389,237, disclose H₂SO₄ enhanced super acidcatalysts; U.S. Pat. Nos. 5,220,095, 5,731,256, 5,489,729, 5,364,976,5,288,685 and European Patent Application No. 714,871A, disclose CF₃SO₃H/silica catalysts; U.S. Pat. Nos. 5,391,527, and 5,739,074, disclosePt—AlCl₃—KCl/Al₂O₃ catalysts; U.S. Pat. Nos. 5,157,196, 5,190,904,5,346,676, 5,221,777, 5,120,897, 5,245,101, 5,012,033, 5,157,197, andpublished PCT Application No. WO 95/126,815, etc. disclose Lewis acidcatalysts, such as SbF₅, BF₃ and AlCl₃; U.S. Pat. Nos. 5,324,881, and5,475,178, disclose supported heteropolyacid catalysts; U.S. Pat. Nos.3,917,738 and 4,384,161, disclose molecular sieve catalysts.Nonetheless, despite continued efforts over 50 years, there is still anunmet need for an improved, stable and economical solid acid alkylationcatalyst.

Solid acid catalysts, such as zeolite catalysts that have a plurality ofH⁺, or acid sites, which are less toxic and less dangerous; however,such catalysts have fewer H⁺, or acid sites than liquid acid catalysts,and only a portion of such acid sites are strong enough to catalyzealkylation reactions. Fundamentally different from liquid acids,zeolites have different populations of sites which differ substantiallyin their nature (Bronsted vs Lewis acids) and strength. Depending on thetype of zeolite, its aluminum content, and the exchange procedure,Brønsted and Lewis acid sites having a wide range of strength andconcentration are present. Zeolites exhibit a considerably lower proton(acid site) concentration than liquid acids. For example, 1 g of H₂SO₄contains 20×10⁻³ moles of protons, whereas 1 g of zeolite HY, with aSi/Al ratio of five, contain no more than 1×10⁻³ moles of protons out ofwhich 20-30% are strong enough to catalyze the alkylation reaction. As aresult, the useful lifetime of a solid-acid catalyst is usually 2 ordersof magnitude shorter than a liquid acid catalyst making it difficult todevelop commercially viable paraffin alkylation technologies usingsolid-acid catalysts.

Methods of making zeolite catalysts having improved characteristics foralkylation have been described by Lercher et al. in U.S. Pat. No.7,459,412. The catalysts described in this patent contain a crystallinezeolite with a silica (SiO₂) to Alumina (Al₂O₃) molar ratio less than10, and an alkali metal content of 0.2 wt % or less. In the examples,Lercher et al. treated a commercial zeolite X with lanthanum nitrate,and then ammonium nitrate, and calcined at 450° C. in flowing air toresult in the low alkali metal content zeolite catalyst. Lercher et al.reported that the catalyst should have the highest possibleconcentration of Bronsted acid centers and a low concentration of strongLewis acid centers. The Lewis acid centers are catalytically inactive,but bind olefins that accelerate oligomerization and deactivation of thecatalyst. Lercher et al. report that the Lewis acid centers arise fromaluminum cations that are released from the crystal lattice during thecalcination step.

Prior art methods that do not combine the rare earth treatments withdeammoniation have described deammoniation temperatures of at least 500°C. See U.S. Pat. Nos. 3,893,942, 3,851,004, and 5,986,158.

The release of aluminum from the zeolite crystal lattice is known asdealumination and occurs at elevated temperature in the presence watervapor. For example, Lutz et al. in “Investigations of the Mechanism ofDealumination of Zeolite Y by Steam: Tuned Mesopore Formation Versus theSi/Al Ratio,” in the Proceedings of the 14^(th) Intl Zeolite Conf., pp.25-30 (2004) reported on the dealumination of zeolite Y at 1 bar watervapor at 500° C., 600° C., and 700° C. showing increasing rates ofdealumination with increasing temperature.

SUMMARY OF THE INVENTION

In a first aspect, the invention provides method of alkylatingisobutane, comprising: under steady state conditions, passing a feedmixture of isobutane and C2-C5 olefins (which is typically conducted ina continuous fashion) into a reaction chamber such that catalyst age is2.5 or greater and producing 5 kg of alkylate product per kg of catalystor greater wherein the olefin conversion remains above 90%, and theResearch Octane Number (RON) of the products remains above 92. Steadystate means that the selectivity to C8 isomers changes by 10% or lessover a time period in which the 5 kg of alkylate product is produced perkg of catalyst. For example, a change in selectivity from 80% to 72%would be a 10% change. In this method, the reaction chamber comprises acrystalline zeolite catalyst; wherein the crystalline zeolite catalystcomprises sodalite cages and supercages, a Si/Al molar ratio of 20 orless, less than 0.5 weight percent alkali metals, and rare earthelements in the range of 10 to 35 wt %. Optionally, the catalyst maycomprise up to 5 wt % Pt and/or Pd; and/or Nickel. The above-mentionedfirst aspect is a subset of a larger aspect having the samecharacteristics except not requiring rare earth elements. Throughout thedescriptions in this specification, percentages of Si, Al, and rareearth elements refer to the elemental composition of the zeolitecrystallites (apart from the binder), which can be easily measuredduring catalyst synthesis, and can be determined spectroscopically or,if necessary, by physical separation of the binder and crystallite inthe finished catalyst. The elemental composition of Pd, Pt, and Ni isbased on the weight percent of the entire particles.

The statement that the catalyst lifetime is 2.5 or greater is notintended to mean that the catalyst age could be infinite, but that themethod operates for sufficient time, without catalyst regeneration forthe catalyst age to be at least 2.5. The method could be operated forsufficient time, without catalyst regeneration for the catalyst age tobe 3.0. In some cases, the method can be described as having a catalystage between 2.5 and 3.5. The method could be operated for a catalyst agegreater than 3.5. A catalyst scientist would understand that thecatalyst synthesis and reaction conditions described here could beoptimized through routine optimization, within the limitations describedin conjunction with the above-described method, to reach a natural limitto the catalyst lifetime.

In various embodiments, the method may further be characterized by oneor any combination of the following options: a reaction temperature ofbetween 45 and 90° C. (in some embodiments between 55 and 80° C., insome embodiments between 60 and 75° C.); an operating pressure of 250 to400 psig; wherein, subsequent to the continuous operation, the catalystis regenerated in a stream of flowing hydrogen at a temperature of atleast 250° C. and a GHSV of 500 l/hr or greater; wherein the feed I/Oratio is 12 or lower, wherein the method is run continuously for 18-36hours at an olefin hourly space velocity of 0.1 l/hr or higher withoutregenerating the catalyst; wherein the reaction chamber comprises apacked catalyst bed; operating the method with a recycle stream suchthat the ratio of the recycle stream flow rate to the feed stream flowrate is 20 or higher; wherein C8 selectivity is at least 70%; whereinthe C2 to C5 olefin consists essentially of butenes; wherein thecatalyst comprises 0.1 wt % to 5 wt % of Pt, Pd, Ni or combinationsthereof; wherein the method is run continuously for a catalyst age of2-3.5 without regenerating the catalyst; comprising a recycle streamsuch that the catalyst be I/O is greater than 300; wherein the C2 to C5olefin consists essentially of mixed butenes; wherein the C2 to C5olefin consists essentially of propylene; wherein the C2 to C5 olefincontains less than 2000 ppm of butadiene; wherein the C2 to C5 olefincontains less than 2 wt % isobutylene; wherein the C2 to C5 olefincontains less than 250 ppm of mercaptans; wherein the C2 to C5 olefincontains less than 300 ppm acetonitrile and less than 200 ppmpriopionitrile; wherein the C2 to C5 olefin contains less than 50 ppmwater; wherein the method is run continuously for a catalyst age of 2.5or a catalyst age of 3.0; and/or wherein the zeolite structure iszeolite X or zeolite Y.

In another aspect, the invention provides a method of making analkylation catalyst, comprising: providing a crystalline zeolitestructure comprising sodalite cages and supercages and having a Si/Almolar ratio of 20 or less, and a first concentration of alkali metal;contacting the zeolite with a solution comprising a rare earth metal;calcining said catalyst by heating said catalyst to a calcinationtemperature of at least 575° C. to produce a catalyst intermediatecomprising the rare earth metal and second concentration of alkali metalthat is less than the first concentration of alkali metal; contactingthe catalyst intermediate with an ammonium solution, drying to removeexcess solution, and then heating the catalyst to generate the hydrogen(active) form of the zeolite—the deammoniation step. In someembodiments, the calcining step comprises heating to at least 575° C.,or at least 600° C., or 575° C. to 625° C., and maintaining thesetemperatures for at least 1 hour or at least 2 hours. The deammoniationstep is typically carried out at least about 400° C. and below 500° C.,in some embodiments in the range of 375 to 425° C. The deammoniationstep is carried out at the stated temperature ranges for at least onehour, preferably at least two hours, or at least 4 hours, in someembodiments in the range of 1 to 10 hours, or 2 to 6 hours. Thedeammoniation step is conducted under the flow of an oxygen-containing,dry gas (typically dry air). The deammoniation step is conductedseparately from the rare earth ion exchange and any subsequentcalcination step. Preferably, the calcining step is conducted in thepresence of a dry gas (preferably dry air); and desirably the entireprocess is conducted in the absence of steam.

The step of contacting the intermediate is separate from and subsequentto the rare earth ion exchange.

The crystalline zeolite structure starting zeolite material may comprise5 to 20 wt %, or 5 to 17 wt %, or 10-17 wt %, or 12-17 wt % Na.

The solution comprising a rare earth metal preferably comprises La³⁺,preferably an aqueous solution comprising 0.5 to 0.8 M La. Experimentswere conducted with La showing the degree of exchange shown below:

Composition of LaX catalyst with varying amounts of Lanthanumconcentration in solution

La+3 Concentration (M) La2O3 (wt %) Degree of Exchange 0.2 19.08 64% 0.419.96 67% 0.6 25.63 85% 0.8 29.18 97%

The “active” form (or hydrogen form as it is sometimes called) is thecatalyst after a deammoniation step that can be used for alkylation.

In some embodiments, the invention includes the alkylation methodsdescribed herein but employing a β-zeolite catalyst in place of the X orY zeolite.

Any of the methods described herein may include a step of regeneratingthe catalyst.

In various preferred embodiments, the method may comprise one or more ofthe following features: wherein the step of calcining to a temperatureof at least 600° C., thereby provides a catalyst in which a portion ofthe alkali metal cation sites are replaced with rare earth metal cationsites and wherein the step of contacting with an ammonium solution,thereby provides a catalyst in which a portion of the alkali metalcation sites are replaced with rare earth metal cation sites, andanother portion of the alkali metal cation sites are replaced withammonium cation sites, and further wherein the deammoniation temperaturedoes not exceed 400° C. in the presence of air, whereby at least aportion of said ammonium cation sites are replaced with H+ sites,thereby providing a catalyst in which a portion of said alkali metalcation sites have been replaced with rare earth metal cation sites andanother portion of said alkali metal cation sites have been replacedwith H+ sites; wherein the rare earth metal is selected from the groupconsisting of lanthanum, cerium, neodymium, and praseodymium, and saidrare earth metal cations are selected from the group consisting oflanthanum cations, cerium cations, neodymium cations, and praseodymiumcations; wherein the rare earth metal comprises lanthanum; wherein thealkali metal cation sites are sodium cation sites; wherein the catalysthas a silica to alumina ratio of from about 2 to about 35; wherein thecatalyst has a silica to alumina ratio of from about 2 to about 10;wherein the solution comprising a rare earth metal comprises an aqueousLa(NO)₃ solution; wherein the solution comprising a rare earth metalcomprises an aqueous La₂(SO₄)₃ solution; wherein the solution comprisinga rare earth metal comprises an aqueous LaCl₃ solution; wherein thesolution comprising a rare earth metal comprises an aqueous solution ofat least 0.1 M Lanthanum ions or at least 0.2 M La, or at least 0.4 MLa, or at least 0.6 M La, or at least 0.8 M La, or in the range of 0.2to 0.8 M La; wherein said catalyst is contacted with the rare earthmetal solution at a temperature of from 60 to 90° C.; wherein thecatalyst is contacted with the rare earth metal solution for a period oftime of about 2 hours; wherein the step of calcining does not exceed600° C.; wherein the step of calcining is conducted from 2 to 8 hours;wherein, during the calcination step, the catalyst is heated in thepresence of air which has a moisture content that does not exceed 2.0wt. % or does not exceed 0.2 wt %; wherein the ammonium solutioncomprises an aqueous solution of at least 0.1 M, or at least 0.2 M, orat least 0.3 M, or at least 0.5 M, or at least 1 M ammonium ions;wherein the step of contacting with an ammonium solution, which providesa catalyst in which a portion of the alkali metal cation sites arereplaced with ammonium cation sites, comprises an aqueous solution ofammonium nitrate or ammonium sulfate; wherein the crystalline zeolite isselected from the group consisting of Zeolite X and Zeolite Y;preferably Zeolite X.

In another aspect, the invention provides an alkylation catalyst,comprising:

a zeolite structure comprising sodalite cages and supercages, a Si/Almolar ratio of 20 or less, less than 0.5 weight percent alkali metals,rare earth elements in the range of 10 to 35 wt % (or a molar ratio ofrare earth elements to (Si and Al) in the range of 0.06 to 0.20); andcharacterizable by a catalyst lifetime greater than 2.5 using a testwhere the solid-acid catalyst is loaded in a fixed-bed reactor such thatthe d_(T)/d_(p)>10 (diameter of tube to diameter of catalyst particles)and L/d_(p)>50 (length of catalyst bed to diameter of catalystparticles) and exposed to a) a feed stream comprising 10:1 molar ratioof isobutane:n-butenes at 60° C. and 300 psig with a recycle ratio of50, where the ratio of system volume to catalyst volume is 7, withoutregeneration, and wherein the RON of the product is at least 92; or b) afeed stream comprising 100:1 molar ratio of isobutane:n-butenes at 60°C. and 300 psig without regeneration, and wherein the RON of the productis at least 92. Optionally, the catalyst may comprise from 0.1 wt % upto 5 wt % Pt and/or Pd; and/or Nickel. Since the catalyst cannot becompletely distinguished from the prior art based solely on itselemental composition, the measurement described above is needed for aunique characterization of the catalyst. In various embodiments, thecatalyst may be further characterized by any of the compositions orphysical characteristics described herein. Preferably, particle size isin the range of 1 to 5 mm. In some preferred embodiments,d_(T)/d_(p)>20, or d_(T)/d_(p)>50 and/or d_(T)/d_(p)>10. Particle sizecan be determined by mesh size and, in preferred embodments, catalystparticle size (including binder) is in the range of 4 mesh to 20 mesh(U.S. mesh size).

In various preferred embodiments, the catalyst may comprise one or moreof the following features: a catalyst lifetime of between 2.5 and 3.5;where the Catalyst Lifetime parameter is defined as the catalyst agewhen the olefin conversion falls below 90% (or, in some preferredembodiments, below 95%) using a test where the solid-acid catalyst isloaded in a fixed-bed reactor such that the dT/dP>10 (diameter of tubeto diameter of catalyst particles) and L/dP>50 (length of catalyst bedto diameter of catalyst particles) and exposed to a flow comprising a) afeed of 10:1 molar ratio of isobutane:n-butenes at 60° C. and 300 psigwith a recycle ratio (R=volumetric flow rate of recyclestream/volumetric flow rate of feed stream) of 50, where V_(S)/V_(C) is7 (the ratio of system volume to catalyst volume), without regeneration,and wherein the RON of the product is at least 92; where the CatalystLifetime parameter is defined as the catalyst age when the olefinconversion falls 90% (or, in some preferred embodiments, below 95%)using a test where the solid-acid catalyst is loaded in a fixed-bedreactor such that the dT/dP>10 (diameter of tube to diameter of catalystparticles) and L/dP>50 (length of catalyst bed to diameter of catalystparticles) and exposed to a flow comprising a feed stream comprising100:1 molar ratio of isobutane:n-butenes at 60° C. and 300 psig withoutregeneration, and wherein the RON of the product is at least 92. Invarious embodiments, the catalyst may have any of the othercharacteristics described anywhere in this specification.

In another aspect, the invention provides a method of making analkylation catalyst, comprising: providing a (a) β-zeolite or (b) acrystalline zeolite structure comprising sodalite cages and supercages;and having a Si/Al molar ratio of 20 or less (preferably 15 or less or10 or less, or in the range of 2.5 to 20 or 15 or 10 or in the range of5 to 20 or 15 or 10), and heating to a temperature in the range of 375to 500° C., preferably 375 to 425° C., or 400 to 450° C., in thepresence of an oxygen-containing gas (typically air) to convert thecatalyst to the active form. In some preferred embodiments, the methodcomprises one or more of the features: the step of heating to atemperature is conducted for at least 1 hour; further comprising a stepof treating the β-zeolite or crystalline zeolite structure with asolution comprising Pd, Pt, or Ni, or a combination thereof; wherein thealkylation catalyst comprises the crystalline zeolite structure, thezeolite structure comprising an alkali metal, and wherein, prior to thestep of heating, contacting the zeolite structure comprising an alkalimetal with an ammonium solution and then drying to remove excesssolution; wherein the active form is disposed in a fixed bed reactor andpassing isobutane and C2 to C5 olefins.

In a further aspect, the invention provides a method of alkylatingisobutane, comprising: passing a feed mixture consisting of excessisobutane and C2 to C5 olefins into a reaction chamber; wherein thereaction chamber comprises a crystalline β-zeolite catalyst; wherein thecrystalline β-zeolite catalyst comprises a Si/Al molar ratio of 20 orless, less than 0.5 weight percent alkali metals; and up to 5 wt % ofPt, Pd and or Nickel, wherein, at steady state, at least 90% of thebutenes (or at least 90% of the C2 to C5 olefins) are converted toproducts and wherein the Research Octane Number (RON) remains above 92;and conducting the process for a catalyst age of 2.5 or greater over thesame catalyst; and wherein steady state means that the selectivity to C8isomers changes by 10% or less over the entire period that the catalystage is determined.

In another aspect, the invention provides an alkylation reaction system,comprising:

a reactor comprising a multi-stage reaction chamber comprising a fixedbed of zeolite catalyst; the reaction chamber comprising 4 to 8 stageseach comprising an inlet, wherein olefin flow is distributed through theinlets into the 4 to 8 stages; a product stream comprising alkylate; anda recycle pump that recycles the product stream at a recycle ratio of 6to 10. In this system, both the solid and fluid elements are consideredcomponents of the system. The term “stages” may alternatively bedescribed as zones. In a preferred embodiment, the reaction chambercomprises 4-6 stages wherein olefin flow is distributed to the 4-6stages.

The invention also includes a β-zeolite catalyst or X or Y zeolitecatalyst, comprising: a Si/Al molar ratio of 20 or less, less than 0.5weight percent alkali metals, and, optionally up to 5 wt % Pt and/or Pd;and/or Nickel; and characterizable by a Catalyst Lifetime of 2 orgreater as described above. The catalysts may include any of thefeatures discussed in this specification.

The invention also includes systems that can be defined as reactorapparatus plus the reactants and/or reactant products and/or conditionswithin the reactor apparatus.

Glossary

Beta-zeolite (β-zeolite)—Is a known form of zeolite that is a highlyintergrown hybrid of two distinct, but closely related structures thatboth have fully three-dimensional pore systems with 12-rings as theminimum constricting apertures. See Newsam et al., “StructuralCharacterization of Zeolite Beta,” Proc. of Royal Soc., A (1988)

Calcination Temperature—The term “calcination temperature” refers to themaximum temperature utilized as an intermediate step in the catalystsynthesis procedure intended to remove the hydration sphere fromlanthanum ions and allow solid-state exchange between lanthanum andsodium cations in the sodalite and supercages.

Deammoniation Temperature—The temperature at which the catalyst isconverted from an ammonium form to its active (H+) form is referred tohere as the “deammoniation temperature.” This step first converts theammonium form of the active sites to Bronsted acid sites (H+), furtherdehydroxylation may convert active sites to Lewis acid sites.

Regeneration Temperature—The solid acid catalyst may be regeneratedunder flowing hydrogen gas at elevated temperatures in order tohydrocrack heavier hydrocarbons and remove them from the zeoliticstructure. The maximum temperature used in this step is referred to asthe “regeneration temperature.”

Conversion—The term “conversion of a reactant” refers to the reactantmole or mass change between a material flowing into a reactor and amaterial flowing out of the reactor divided by the moles or mass ofreactant in the material flowing into the reactor. For example, if 100grams of olefin are fed to a reactor and 10 grams of olefin exit thereactor, the conversion is [(100−10)/100]=90% conversion of olefin.

Olefins—As used herein, the terms “olefin” or “olefin compound” (a.k.a.“alkenes”) are given their ordinary meaning in the art, and are used torefer to any unsaturated hydrocarbon containing one or more pairs ofcarbon atoms linked by a double bond. In this invention, C₂-C₅ olefinsrefers to ethylene, propylene, n-butylenes, isobutylene, and the variousisomers of pentene. The phrase “C₂-C₅ olefins” has the standard meaningencompassing any combination of olefins in the C2 to C5 range, with nominimum requirement for any of the C2 to C5 compounds. In some preferredembodiments, the C2 to C5 olefins contain at least 50 wt % C4 olefins asa percent of all olefins.

Olefin Hourly Space Velocity (OHSV)—“Olefin Hourly Space Velocity” isdefined as the mass flow rate of olefin fed to the reactor per hourdivided by the mass of catalyst in the reactor.

Recycle Ratio (R)—The “recycle ratio” is defined as the volumetric flowrate of the recycle stream divided by the volumetric flow rate of thefeed stream.

Feed I/O Ratio—In alkylation experiments the concentration of olefin inthe feed is often expressed as a ratio of isobutane to olefinic speciesby mole in the reactor feed (“Feed I/O Ratio”). A Feed I/O Ratio(I/O_(feed)) of 100 is equivalent to ˜1% olefin in isobutane—I/O_(feed)of 10 is equivalent to ˜9.1% olefin in isobutane.

Catalyst Bed I/O Ratio—In the reactor schemes discussed in this patent,the solid acid catalyst is exposed to isobutane to olefin ratios muchhigher than those in the reactor feed. This is achieved by bothspreading the feed over a number of evenly distributed inlets along thecatalyst bed (N=the number of inlets) and/or utilizing a recycle stream.The “catalyst bed I/O ratio” is calculated as:I/O_(bed)=N*{I/O_(feed)+R(I/O_(feed)−1)).

System Volume—The total volume of the reactor and any tubing for feeddelivery, recycle lines, sample extraction from the exit of the feedpumps to the back pressure regulator of the reactor system.

Catalyst volume: The bulk volume where solid catalyst (including voidspaces between catalyst particles) is present at reaction conditions isdefined as the “catalyst volume.” In the preferred case of a fixedcatalyst bed, the bulk volume includes catalyst particles and voidsbetween particles. For example, in one preferred embodiment, a fixedcatalyst has 0.62 catalyst particle fraction and 0.38 void fraction. Insome embodiments of the invention, the catalyst bed comprises between0.38 to 0.85 void fraction.

Catalyst Age—“Catalyst age” is the mass of olefin fed to the reactordivided by the mass of catalyst.

Catalyst Lifetime—The catalyst age at which the olefin conversion fallsbelow 90% is defined as the “catalyst lifetime.”

Pore size—Pore size relates to the size of a molecule or atom that canpenetrate into the pores of a material. As used herein, the term “poresize” for zeolites and similar catalyst compositions refers to theNorman radii adjusted pore size well known to those skilled in the art.Determination of Norman radii adjusted pore size is described, forexample, in Cook, M.; Conner, W. C., “How big are the pores ofzeolites?” Proceedings of the International Zeolite Conference, 12th,Baltimore, Jul. 5-10, 1998; (1999), 1, pp 409-414.

One of ordinary skill in the art will understand how to determine thepore size (e.g., minimum pore size, average of minimum pore sizes) in acatalyst. For example, x-ray diffraction (XRD) can be used to determineatomic coordinates. XRD techniques for the determination of pore sizeare described, for example, in Pecharsky, V. K. et at, “Fundamentals ofPowder Diffraction and Structural Characterization of Materials,”Springer Science+Business Media, Inc., New York, 2005. Other techniquesthat may be useful in determining pore sizes (e.g., zeolite pore sizes)include, for example, helium pycnometry or low-pressure argon adsorptiontechniques. These and other techniques are described in Magee, J. S. etat, “Fluid Catalytic Cracking: Science and Technology,” ElsevierPublishing Company, Jul. 1, 1993, pp. 185-195. Pore sizes of mesoporouscatalysts may be determined using, for example, nitrogen adsorptiontechniques, as described in Gregg, S. J. at al, “Adsorption, SurfaceArea and Porosity,” 2nd Ed., Academic Press Inc., New York, 1982 andRouquerol, F. et al, “Adsorption by powders and porous materials.Principles, Methodology and Applications,” Academic Press Inc., NewYork, 1998.

Residence Time—Residence time is the time a substance is in the reactionvessel. It can be defined as the volume of the reactor divided by theflow rate (by volume per second) of gases into the reactor.

Selectivity—The term “selectivity” refers to the amount of production ofa particular product (or products) as a percent of all productsresulting from a reaction. For example, if 100 grams of products areproduced in a reaction and 80 grams of octane are found in theseproducts, the selectivity to octane amongst all products is 80/100=80%.Selectivity can be calculated on a mass basis, as in the aforementionedexample, or it can be calculated on a molar basis, where the selectivityis calculated by dividing the moles a particular product by the moles ofall products. Unless specified otherwise, selectivity is on a massbasis.

Yield—The term “yield” is used herein to refer to the amount of aproduct flowing out of a reactor divided by the amount of reactantflowing into the reactor, usually expressed as a percentage or fraction.Mass yield is the mass of a particular product divided by the weight offeed used to prepare that product.

When unspecified, “%” refers to mass % which is synonymous with weight%. Ideal gas behavior is assumed so that mole % is the same as volume %in the gas phase.

As is standard patent terminology, the term “comprising” means“including” and does not exclude additional components. Any of theinventive aspects described in conjunction with the term “comprising”also include narrower embodiments in which the term “comprising” isreplaced by the narrower terms “consisting essentially of” or“consisting of.” As used in this specification, the terms “includes” or“including” should not be read as limiting the invention but, rather,listing exemplary components. As is standard terminology, “systems”include to apparatus and materials (such as reactants and products) andconditions within the apparatus.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1: Faujasite Structure of zeolite X and Y

FIG. 2: Fixed-Bed Reactor configurations suitable for solid-acidalkylation

FIG. 3: Fixed-bed micro-reactor for catalyst tests

FIG. 4: Detailed reactor dimensions for catalyst tests. L is the lengthof the catalyst bed, dT is the tube diameter and dP is the catalystparticle diameter.

FIG. 5: Representative solid-acid alkylation experimental results forCatalyst A.

FIG. 6: Catalyst lifetime for lanthanum exchanged X zeolite calcined atdifferent temperatures (catalysts A-D).

FIG. 7: Catalyst lifetime for Y zeolite with (E, top line) and without(G, bottom line) a lanthanum exchange and subsequent calcination.

FIG. 8: Catalyst lifetime for lanthanum exchanged X zeolite deammoniatedwith different moisture levels of <2 ppm (catalyst D) and 1.2% by vol.(catalyst H).

FIG. 9: Catalyst lifetime for lanthanum exchanged X zeolite deammoniatedat different temperatures (catalysts D and I-L).

FIG. 10: Catalyst lifetime for lanthanum exchanged X zeolite at varyingLa+3 concentrations.

FIG. 11: Catalyst lifetime for β zeolite deammoniated at differenttemperatures (catalysts M-P).

FIG. 12: Single stage feed recycle reactor schematic.

FIG. 13: Representative recycle solid-acid alkylation experimentalresults for Catalyst D.

FIG. 14: Representative Octane numbers for recycle solid-acid alkylationusing Catalyst D.

FIG. 15: Long term stability data results for Catalyst W.

FIG. 16: Long term stability data results for Catalyst X.

FIG. 17 shows catalyst lifetime as a function of deammoniationtemperature and SAR for β zeolite catalysts.

FIG. 18. Butene conversion and product octane as a function of catalystage for the Y zeolite deammoniated at 400° C. (Catalyst F).

DETAILED DESCRIPTION OF THE INVENTION

The invention relies on a solid acid, crystalline zeolite structure thathas both supercages and sodalite cages. These structures are well knownand are shown in FIG. 1. It is believed that Bronsted acid sites in thesupercages are important in catalyzing the alkylation reaction. TheFaujasite structure with supercages and sodalite cages is known inZeolite X and zeolite Y. Commercial zeolites are typically mixed withbinders. The purity of zeolites can be measured by oxygen adsorption asdescribed in Experimental Methods in Catalytic Research, vol. 2,Anderson and Dawson eds. (1976). Thus, although the invention canoperate with pure crystalline zeolite, it typically operates withcrystalline zeolite mixed with binder or other materials. Relativelyhigh levels of aluminum are desirable; thus it is desirable to usezeolites with a Si/Al molar ratio of 20 or less, preferably 15 or less,in some embodiments in the range of 1 to 12, in some embodiments in therange of 2 to 10. We have found excellent results based on Zeolite X andacceptable performance with zeolite Y; nonetheless, based on thesimilarity of the pore structure, it is believed that zeolite Y shouldalso function as a suitable starting material for the catalyst of theinvention. The pore size of the sodalite cage structure does not exceedabout 8 Angstroms, and the supercage structure has pores in which thepore size is at least about 10 Angstroms. In some preferred embodiments,the Si/Al molar ratio of zeolite X is in the range of 1 to 2; the Si/Almolar ratio of zeolite Y is in the range of 2 to 5; the Si/Al molarratio of beta zeolite is in the range of 8 to 13.

In a method of making a catalyst according to the present invention, amaterial containing a crystalline zeolite structure comprising sodalitecages and supercages and having a Si/Al molar ratio of 20 or less istreated with a solution containing a rare earth metal. The crystallinezeolite structure contains an alkali metal, typically sodium orpotassium, most typically, sodium. The amount of alkali metal in thestarting material is typically greater than 1 wt %, in some preferredembodiments greater than 3 wt %, in some embodiments between 5 and 20 wt%. The solution containing a rare earth metal is typically an aqueoussolution. Preferred rare earth metals comprise lanthanum, cerium,neodymium, and praseodymium, and mixtures thereof; most preferablycomprise lanthanum (La), and in some preferred embodiments the rareearth metal is at least 90% La or at least 95% La (by weight relative tototal weight of all rare earth metals in solution). Preferably, thezeolite is treated with the rare earth solution at elevated temperature,preferably from 60 to 95° C., more preferably 70 to 90° C.; typicallywith a nitrate or sulfate salt solution. The solution containing a rareearth metal preferably has a concentration in the range of 0.1 M to 1.0M, in some embodiments in the range of 0.4 to 0.8 M. Multipletreatments, for example, 3 treatments are preferred. Each treatment ispreferably conducted for at least one hour at the elevated temperature,in some embodiments between 1 and 4 hours.

If there is excess solution containing rare earth metal, it can beremoved by decanting or filtering. Optionally, after decanting orfiltering, the treated zeolite can be dried at temperatures up to 100°C. The resulting material is believed to have rare earth metal locatedin the supercages, but not yet exchanged with the alkali metal in thesodalite cages.

To effectuate exchange of alkali ions in the sodalite cages with therare earth ions located in the supercages, the catalyst is calcined at atemperature of at least 575° C. Although it was reported that the amountof La⁺³ in the sodalite cages becomes constant at temperatures above300° C. (Monsalve, Thesis “Active Acid Sites in Zeolite CatalyzedIso-butane/cis-2-butene Alkylation” Chap. 3, p 4), we surprisingly foundsignificantly improved results from calcining at a much highertemperature. Preferably, the calcining step is carried out at atemperature of 575 to 650° C. In some preferred embodiments, the zeoliteis held at a temperature between about 90 and 110° C. for at least 10minutes, preferably at least 50 minutes. The zeolite can be heated atany suitable temperature ramping rate; for example between 1° C./min to10° C./min. It may be preferred to hold the temperature at anintermediate value, such as between 200 and 300° C. for 30 min or more.Preferably, the zeolite is maintained at a temperature of at least 575°C., preferably between 575 and 650° C., in some embodiments between 600and 625° C., or from 575 to 600° C., for at least 50 minutes, preferablyfor at least about 100 minutes; in some embodiments for between 50 and500 minutes, in some embodiments between 50 and 240 minutes. Preferably,the entire calcination process, including temperature ramping times, iscompleted within 1 day, or completed within 2 days. The calcination stepis preferably carried out at a relatively low humidity, for example, indry flowing air containing less than 1 mass % water, in some embodimentsless than about 50 ppm water. We believe that the calcination stepcauses some and, preferably essentially all, of the alkali metal ions(usually Na⁺) in the sodalite cages to be replaced with the rare earthions (preferably La+3) from the sodalite cages.

After calcination, the calcined zeolite is cooled and treated with anammonium solution. The solution preferably has an ammonia concentrationin the range of 0.1 M to 1.0 M, in some embodiments in the range of 0.2to 0.5 M. This can be repeated several times; for example, from 2 to 5times. One preferred set of conditions for the ammonium treatment is atemperature of from 50 to 100° C. for 10 minutes to 4 hours or more;more preferably from 30 minutes to two hours. In some embodiments of theinvention, there is no rare earth exchange step and the zeolite(typically zeolite Y; containing Na cations) can be treated by theammoniation process described herein.

Any excess solution can be removed by decanting or filtration. Theammonium-exchanged zeolite can be heated to drive off excess water, forexample to 100° C. or 200° C.

Prior to use as a catalyst, the zeolite is converted from its ammoniumform to the hydrogen form by heating, preferably in an atmosphere havingvery little water; for example, 1 mass % or less, or 0.2 mass %, or 2ppm or less of water. This deammoniation temperature is preferably inthe range of 300 to 400° C., more preferably 350 to 400° C.

Although the scope of the present invention is not to be limited to anytheoretical reasoning, it is believed that the deammoniation stepconverts the ammonium cation sites to Bronsted acid sites, especially inthe supercages, while the rare earth elements remain in the sodalitecages. Because the acid, or H+, sites are located in the larger diametersupercage structure of the catalyst, pore mouth plugging issignificantly reduced, allowing the catalyst to remain active forincreased periods of time, while the rare earth metal cation sites, suchas, for example, La⁺³ sites, provide enhanced stability to the sodalitestructure. We believe that at least 80% of the cationic sites in thesodalite portion are rare earth metal cation sites, and at least 80% ofthe cationic sites in the supercage portion are H+ sites.

We have found that careful control of the deammoniation conditions forthe zeolite catalyst lead to improvements in catalyst performance, whenconverting the ammonium form of the zeolite to the active or acid form.When the ammonium form of a zeolite is heated, the initial step is theevolution of physically adsorbed water, which causes a first endothermat about 150° C.; this step is completed at 200° C. Ammonia then isevolved which gives rise to a second endotherm at 300° C.; this step iscompleted at about 400° C. Raising the temperature above 400° C. resultsin evolution of water from the condensation of the hydroxyl groups. Thisdehydroxylation step results in a) a significant decrease in the numberof active catalytic acid sites and b) conversion of the preferredBronsted acid sites to the Lewis acid sites which increases the rate ofcatalyst deactivation.

The invention also relates to a reactor suitable for paraffin alkylationusing solid acid catalysts. Paraffin alkylation is a fast reaction,which benefits from low olefin concentrations (typically the reactor I/Oratio>300) in the reactor to suppress the polymerization reaction. Inconventional liquid-acid based reactors, high speed mechanical agitatorsare used to disperse the hydrocarbon feed into the acidic medium.Specially designed jets are used to introduce the olefin feed as smalldroplets to avoid high localized olefin concentration. A departure fromperfect mixing conditions results in significant deterioration ofproduct octane quality and formation of Acid Soluble Oils via olefinpolymerization reaction which leads to higher acid consumption. The onlyway to achieve the same level of mixing with solid-catalysts, is to usea slurry system. However, slurry systems are difficult to handle andequipment needed to pump slurries around are very expensive.

Fixed-bed reactors are easier to design, scale-up and maintain and,therefore, preferred embodiments utilize a fixed bed reactor. One way ofachieving a low olefin concentration in the bulk liquid is obtained bystaging the olefin feed over the catalyst bed. This approach is oftenused in designing reactors for aromatic alkylation reactions for theproduction of ethylbenzene or cumene. Typically 4-6 stages (FIG. 2A) areused in designing such reactors. In paraffin alkylation reactions, thefeed mixture of iso-butane to olefins (I/O) range from 10-15 molar and atypical catalyst bed I/O target is 300-500. This implies that the olefinmust be distributed in 30-50-stages to meet this recommended dilution inolefin concentration which is difficult to achieve in commercial scalereactors. Another way of achieving this level of olefin dilution withoutresorting to staging the feed is by introducing a recycle stream of thereactor effluent (FIG. 2B). By using a recycle ratio of 30-50, a feedwith an I/O ratio of 10-15 would produce a catalyst bed I/Oof 300-500.However, the large volumetric flow through the catalyst bed wouldgenerate a proportionally large pressure-drop greatly increasing theduty (hence power consumption) on the recycle pump. A novel way ofdesigning a fixed-bed reactor for solid acid catalyzed reactions is tocombine the two concepts into a hybrid reactor (FIG. 2C) where theolefin feed is staged over 4-6 zones and the recycle pump recycles theproduct stream at a modest recycle ratio of 6-10. This allows a catalystbed I/O of 300-500 without resorting to a large number of feed stages ora very high recycle ratio. An added benefit of this design is theability to remove the heat of reaction externally. The inventionincludes methods using the reactors described herein and includessystems (apparatus plus fluids and, optionally, conditions within theapparatus) that includes the reactors described herein.

The invention is further elucidated in the examples below. In somepreferred embodiments, the invention may be further characterized by anyselected descriptions from the examples, for example, within ±20% (orwithin ±10%) of any of the values in any of the examples, tables orfigures; however, the scope of the present invention, in its broaderaspects, is not intended to be limited by these examples.

EXAMPLES Example 1 Catalyst A

The starting material was a commercial zeolite X having a SiO₂/Al₂O₃molar ratio of 2.8 (Si/Al of 1.4) and a sodium content of 15% by weight.5 grams of the zeolite was crushed and sieved to 0.5-1.4 mm particles.They were suspended in 50 mL of deionized water and stirred for 15minutes after which the water was decanted. This washing procedure wasrepeated a second time.

A lanthanum ion exchange was performed immediately following the initialwater wash. The zeolite was suspended in 50 mL of a 0.8 M lanthanumnitrate solution and heated to 80° C. while stirring for 2 hours. Thelanthanum solution was decanted and replaced with a fresh solution. Thislanthanum exchange was performed three times followed by 2 water washesof 75 mL each. The zeolite was then left to dry at room temperature.

Following the lanthanum exchange, the zeolite was calcined in a mufflefurnace. The temperature program for calcination was 1.5° C./min ramp to100° C. where it was held for 1 hour, 2.0° C./min ramp to 230° C. andhold for 2 hours, 10° C./min ramp to the final calcination temperatureof 400° C. for 4 hours.

The lanthanum exchanged zeolite was suspended in a 0.5 M ammoniumnitrate solution and heated to 80° C. with stirring for 2 hours. Theammonium solution was decanted and replaced with fresh solution. Thision exchange was performed 3 times followed by 2 water washes of 75 mLeach. The zeolite was then left to dry at room temperature. The zeolitewas deammoniated in dry air (<2 ppm) using the following temperatureprogram: 100° C. (0.5 hours), 120° C. (1 hour), 230° C. (2 hours), 400°C. (4 hours). 400° C. is the deammoniation temperature required toconvert the catalyst from the ammonium form to the active proton form.The lower temperatures are necessary to completely dry the catalyst.

Example 2 Catalyst B

The catalyst was prepared as in Example 1 with the only difference beingthe final calcination temperature. In this example the final calcinationtemperature following lanthanum exchange was 450° C.

Example 3 Catalyst C

The catalyst was prepared as in Example 1 with the only difference beingthe final calcination temperature. In this example the final calcinationtemperature following lanthanum exchange was 550° C.

Example 4 Catalyst D

The catalyst was prepared as in Example 1 with the only difference beingthe final calcination temperature. In this example the final calcinationtemperature following lanthanum exchange was 600° C.

Example 5 Catalyst E

The catalyst was prepared as in Example 1. However, the startingmaterial used was a Y zeolite in this example. The commercial Y zeolitehad a SiO₂/Al₂O₃ molar ratio of 5.0 and a sodium content of 14% byweight. Since the Y zeolite is in powder form it must be filtered ratherthan decanted in each solution exchange. Additionally, it is pelletizedfollowing ammonium exchange and drying then crushed and sieved to0.5-1.4 mm catalyst particles.

Example 6 Catalyst F

The catalyst was prepared as in Example 5 with the only difference beingthat no Lanthanum exchange and subsequent calcination was performed.Following the initial water wash, the Y zeolite undergoes an ammoniumexchange and deammoniation. In this example the deammoniationtemperature was 400° C.

Example 7 Catalyst G

The catalyst was prepared as in Example 5 with the only difference beingthat no Lanthanum exchange and subsequent calcination was performed.Following the initial water wash, the Y zeolite undergoes an ammoniumexchange and deammoniation. In this example the deammoniationtemperature was 550° C.

Example 8 Catalyst H

The catalyst was prepared as in Example 3 with the only difference beingwater content of the air used for activation following ammoniumexchange. In this example, the water content was 1.2% by volume.

Example 9 Catalyst I

The catalyst was prepared as in Example 3 with the only difference beingthe deammoniation temperature used following ammonium exchange. In thisexample the deammoniation temperature was 300° C.

Example 10 Catalyst J

The catalyst was prepared as in Example 3 with the only difference beingthe deammoniation temperature used following ammonium exchange. In thisexample the activation temperature was 350° C.

Example 11 Catalyst K

The catalyst was prepared as in Example 3 with the only difference beingthe deammoniation temperature used following ammonium exchange. In thisexample the deammoniation temperature was 450° C.

Example 12 Catalyst L

The catalyst was prepared as in Example 3 with the only difference beingthe deammoniation temperature used following ammonium exchange. In thisexample the deammoniation temperature was 500° C.

Example 13 Catalyst M

The catalyst was prepared as in Example 1. However, the startingmaterial used was a β zeolite in this example. The commercial β zeolitehad a SiO₂/Al₂O₃ molar ratio of 16. The β zeolite does not undergo alanthanum exchange and the subsequent calcination. Following an initialwater wash, it is immediately exchanged with ammonium 3 times. It isthen deammoniated in dry air with a final temperature of 400° C.

Example 14 Catalyst N

The catalyst was prepared as in Example 13 with the only differencebeing the deammoniation temperature used following ammonium exchange. Inthis example the deammoniation temperature was 450° C.

Example 15 Catalyst O

The catalyst was prepared as in Example 13 with the only differencebeing the deammoniation temperature used following ammonium exchange. Inthis example, the deammoniation temperature was 500° C.

Example 16 Catalyst P

The catalyst was prepared as in Example 13 with the only differencebeing the deammoniation temperature used following ammonium exchange. Inthis example the deammoniation temperature was 550° C.

Example 17 Catalyst Q

The catalyst was prepared as in Example 13 with the only differencebeing the starting β-zeolite had a SiO₂/Al₂O₃ molar ratio of 25.

Example 18 Catalyst R

The catalyst was prepared as in Example 13 with the only differencebeing the starting β-zeolite had a SiO₂/Al₂O₃ molar ratio of 75.

Example 19 Catalyst S

The catalyst was prepared as in Example 17 with the only differencebeing the deammoniation temperature used following ammonium exchange. Inthis example the deammoniation temperature was 550° C.

Example 20 Catalyst T

The catalyst was prepared as in Example 18 with the only differencebeing the deammoniation temperature used following ammonium exchange. Inthis example the deammoniation temperature was 550° C.

Example 21 Catalyst U

The catalyst was prepared as in Example 4 with the only difference beingthe lanthanum ion exchange was performed using a 0.3 M lanthanum nitratesolution.

Example 22 Catalyst V

The catalyst was prepared as in Example 4 with the only difference beingthe lanthanum ion exchange was performed using a 0.5 M lanthanum nitratesolution.

Example 23 Catalyst W

The catalyst was prepared as in Example 4 with the only difference beingthe lanthanum ion exchange was performed using a 0.6 M lanthanum nitratesolution.

Example 24 Catalyst X

The catalyst was prepared as in Example 4 with the only difference beingthe lanthanum ion exchange was performed using a 0.8 M lanthanum nitratesolution.

Example 25 Catalyst Y

The catalyst was prepared as in Example 4 with the only difference beingthe lanthanum ion exchange was performed using a 1.0 M lanthanum nitratesolution.

Example 26 Catalyst Z

The catalyst was prepared as in Example 21. The catalyst was impregnatedwith Tetraamine Platinum Chloride to give 0.1 wt % Pt loading on thecatalyst.

Example 27 Catalyst AA

The catalyst was prepared as in Example 24. The catalyst was impregnatedwith Nickel Nitrate to give 0.25 wt % Nickel loading on the catalyst.

Alkylation activity experiments were performed using an isothermalpacked bed reactor setup. Heating is controlled using an Omegatemperature control unit and a ceramic heating element. Feeds are sentthrough a preheater of ˜75 cm length prior to entering the reactor.

The catalyst of interest (1 g) is first loaded into a reactor shown inFIG. 3 (7.9 mm diameter), a center thermocouple (K-type) is inserted andpositioned such that the tip of the thermocouple (3.1 mm diameter) is atthe bottom of the catalyst bed. 1 mm glass beads are used to fill anyvoid space in the reactor. The catalyst is deammoniated in dry air(GHSV=1000 hr⁻¹) at atmospheric pressure using the following temperatureprogram: 100° C. (0.5 hour), 120° C. (1 hour), 230° C. (2 hours), 400°C. (4 hours) (these values are for Example 1). Following deammoniationthe reactor is allowed to cool to reaction temperature (75° C.), thenpurged with dry nitrogen (GHSV=1000 hr⁻¹) for 0.5 hours. The reactor ispressurized (300 psig) with pure isobutane to begin the experiment.

The reaction feed is contained in helium-purged Hoke cylinders.Isobutane and 1-butene (source for both is AGL Welding Supply Co, Ltd)are analyzed for any impurities using a HP5890 GC equipped with aPetrocol DH column. All feed and product analysis uses this GC systemwith the following program: 60° C. (16 min), ramp at 15° C./min to 245°C. and soak (20 min).

The experiment is run using an olefin hourly space velocity equal to 0.5hr⁻¹ and a feed I/O ratio of ˜100. This equates to 40 g/hr feed rate forisobutane and 0.4 g/hr for 1-butene. The flow rates are controlled byEldex ReciPro Model A pumps. Product samples are extracted using a highpressure sampling port and syringe (Vici Precision Sampling) andimmediately injected into the HP5890 GC for analysis.

Regeneration may be performed using hydrogen gas (1000 hr⁻¹ GHSV) at aregeneration temperature of 250° C. for 2 hours. Process and detailedreactor schematics are shown in FIGS. 3 and 4.

Application Example 1

The lanthanum exchanged X zeolites were prepared with differentcalcination temperatures as in Examples 1-4 (Catalyst A-D). 1 gram ofeach catalyst was loaded into a reactor shown in FIG. 3. The reactor waspurged with nitrogen at a temperature of 75° C. It is then pressurizedwith isobutane to 300 psig. The reaction feed mixture, I/O Ratio of 100,was fed to the reactor with an OHSV of 0.5 hr⁻¹. Product samples werewithdrawn periodically from a high pressure sample port and analyzedusing a gas chromatograph equipped with a Petrocol DH 100 m column.

FIG. 5 shows representative data using catalyst A from Example 1.Research Octane Number (RON) and Motor Octane Number (MON) arecalculated for the alkylate from product analysis.

FIG. 6 shows catalyst lifetime as a function of calcination temperature.The zeolite calcined at 600° C. (Catalyst D) possessed a lifetime of3.75.

As can be seen from FIG. 6, catalyst lifetime plateaus between 450 and550° C.—an observation which matches the conventional understanding thatthe zeolite catalyst be calcined at a temperature of 450° C. or less.Surprisingly, however, we discovered that heating to 575° C. or higher,more preferably 600° C. resulted in a substantial improvement incatalyst lifetime, and preferably less than 650° C. to avoid degradationof the zeolite structure.

Application Example 2

The Y zeolites were prepared with and without lanthanum exchange stepsfollowed by calcination as in Examples 5 and 7 (Catalysts E-G). Theexperimental conditions are identical to those of Application Example 1.

FIG. 7 shows butene conversion as a function of catalyst age for the Yzeolites. As shown, the catalyst age of the H form of zeolite Y issubstantially increased by lanthanum exchange followed by calcination.

Application Example 3

The catalysts used were catalyst D (<2 ppm) and catalyst H (Example 8,1.2% by volume) at different water contents in the air duringdeammoniation. The experiment is identical to Application Example 1.

FIG. 8 shows catalyst lifetime for both the water concentrations usedduring deammoniation. Moisture in the oxidizing gas (typically air)during deammoniation should be avoided. Therefore, the atmosphere (or,more typically, the gas that flows over the zeolite as measured prior tocontacting the zeolite) during deammoniation preferably comprises 0.2vol % or less, more preferably 10 ppm or less, more preferably 5 ppm orless, and still more preferably 2 ppm or less water.

Application Example 4

The catalysts used were from examples 4 (catalyst D) and 9-12 (catalystsI-L). They were deammoniated at different temperatures under dryconditions (<2 ppm). FIG. 9 compares lifetime for the catalystsdeammoniated at different temperatures. There is a maximum catalystlifetime of 2.9 at 400° C. deammoniation temperature. The deammoniationtemperature was maintained at 400° C. for 4 hours.

The superior catalyst lifetime results for deammoniation in the range ofabout 400 to 450° C. was especially surprising since the guidelines fromLinde Molecular Sieves—“Catalyst Bulletin, Ion-Exchange and MetalLoading Procedures” state that to decationize NH₄ ⁺ exchanged molecularsieve should be conducted in dry air at 550° C. for 3-4 hours.

Application Example 5

The catalysts used were from examples 18-22 (catalysts R-V). They weredeammoniated at 400° C. under dry conditions (<2 ppm). FIG. 10 compareslifetime for the catalysts ion-exchanged with varying concentrations ofLanthanum Nitrate solutions. There is a maximum catalyst lifetime of 3.7at 0.8 M Lanthanum Nitrate concentration. FIG. 10 shows that catalystlifetime reaches a maximum, in a single exchange step, where Laconcentration is in the range of about 0.5 to 0.9 M, preferably 0.6 to0.8 M; La concentrations above this range lowers pH and thus causesstructural collapse.

Application Example 6

The β zeolites were prepared with different deammoniation temperaturesas in Examples 13-16 (Catalyst M-P) and loaded into a fixed-bed reactor.In this experiment the reaction was run in recycle mode. The reactionfeed mixture, I/O Ratio of 15, was fed to the reactor at a rate of 10g/hr. The recycle stream flow rate was 40 g/hr. The combined feed rateto the reactor was 50 g/hr with an OHSV of 0.2 hr^(˜1). Product sampleswere withdrawn periodically from a high pressure sample port andanalyzed using a gas chromatograph equipped with a Petrocol DH 100 mcolumn as in Application Example 1

FIG. 11 shows catalyst lifetime as a function of deammoniationtemperature for these β catalysts. The highest performance is for the βcatalyst deammoniated at 400° C. with dry air. Thus, we observed thesurprising result that the deammoniation temperature (450° C. to 400 C,preferably 425° C. to 400 C) resulted in superior catalyst lifetimes.This was a very surprising result since it had been reported thatactivation at 450 C under dry conditions resulted in “barely activecatalysts” and that a temperature of 550 C was required forsignificantly improved activity. See Kunkeler et al., “Zeolite Beta: TheRelationship between Calcination Procedure, Aluminum Configuration, andLewis Acidity”

Application Example 7

The lanthanum exchanged X zeolite from Example 4 (Catalyst D) was loadedinto a fixed-bed reactor with product recycle shown in FIG. 12. Areaction feed mixture, with a I/O molar ratio of 10 (with n-butene asolefin), was fed to a bench-scale reactor with an OHSV of 0.1 hr⁻¹ at atemperature of 60° C. and a pressure of 300 psig. This recycle flow ratewas established such that the recycle ratio was 50. Product samples werewithdrawn periodically from a high pressure sample port and analyzedusing a gas chromatograph equipped with a Petrocol DH 100 m column.

FIG. 13 shows representative product distribution data using catalyst Dfrom Example 4.

FIG. 14 shows reaction octane numbers as a function of run time.Research Octane Number (RON) and Motor Octane Number (MON) arecalculated for the alkylate from product analysis.

The lifetime of this catalyst was >3.25 under commercial reactionconditions before regeneration. The steady state product C₈ selectivitywas 79 wt %, RON was 97 and the product MON was 93.

Application Example 8

The lanthanum exchanged X zeolite from Example 23 (Catalyst Z) wasloaded into a fixed-bed reactor with product recycle shown in FIG. 15. Areaction feed mixture, with a I/O molar ratio of 10 (with MTBE raffinateas olefin), was fed to a bench-scale reactor with an OHSV of 0.1 hr⁻¹ ata temperature of 45° C. and a pressure of 300 psig. The catalyst testwas run for 24 hours and then regenerated with hydrogen gas at 250° C.for 2 hours. This cycle was repeated for 18 months. Results of this testare shown in FIG. 15.

Data shown in FIG. 15 demonstrates that 0.1 wt % Platinum loading on thecatalyst is adequate for regeneration the catalyst with hydrogen gas.

Application Example 9

The lanthanum exchanged X zeolite from Example 27 (Catalyst AA) wasloaded into a fixed-bed reactor with product recycle shown in FIG. 16. Areaction feed mixture, with a I/O molar ratio of 10 (with MTBE raffinateas olefin), was fed to a bench-scale reactor with an OHSV of 0.1 hr-1 ata temperature of 45° C. and a pressure of 300 psig. The catalyst testwas run for 24 hours and then regenerated with hydrogen gas at 250° C.for 2 hours. This cycle was repeated for 50 days.

The data shown in FIG. 16 demonstrates that about 0.25 wt % Nickelloading on the catalyst is adequate for regeneration the catalyst withhydrogen gas; a preferred range is 0.1 to 0.5 wt % Ni.

Application Example 10

The β zeolites were prepared with different Silica-to-Alumina Ratios(SAR) and deammoniation temperatures as in Examples 17-20 (CatalystsQ-T) and loaded into a fixed-bed reactor with product recycle shown inFIG. 14. The reaction feed mixture, I/O Ratio of 15, was fed to thereactor at a rate of 10 g/hr. The recycle stream flow rate was 40 g/hr.The combined feed rate to the reactor was 50 g/hr with an OHSV of 0.2hr-1. Product samples were withdrawn periodically from a high pressuresample port and analyzed using a gas chromatograph equipped with aPetrocol DH 100 m column as in Application Example 1.

FIG. 17 shows catalyst lifetime as a function of deammoniationtemperature and SAR for these β catalysts. The highest performance isfor the β catalyst at SAR 16 and deammoniated at 400° C. Results shownin FIG. 17 are unique in the way the NH4+ form is converted to the H+form. The typical deammoniation temperature used for zeolites istypically 550° C. or higher. Our results clearly show a much superiorperformance at lower deammoniation temperatures.

Application Example 11

The Y zeolites were prepared without Lanthanum exchange steps followedby deammoniation as in Examples 6 (Catalysts F). The experimentalconditions are identical to those of Application Example 1

FIG. 18 shows butene conversion and product octane as a function ofcatalyst age for the Y zeolite deammoniated at 400° C. (CatalystF)—showing superior catalyst lifetime for deammoniation at thistemperature.

Comparing performance of Y-zeolite deammoniated at 400° C. (Catalyst F)with Y-zeolite deammoniated at 550° C. (catalyst G) clearly demonstratesthe superiority of the low temperature deammoniation method.

It is to be understood, however, that the scope of the present inventionis not to be limited to the specific embodiments described above. Theinvention may be practiced other than as particularly described andstill be within the scope of the accompanying claims.

What is claimed is:
 1. A method of alkylating isobutane, comprising:passing a feed mixture consisting of excess isobutane in admixture withbutenes (or C2 to C5) olefins into a reaction chamber; wherein thereaction chamber comprises a crystalline zeolite catalyst; wherein thecrystalline zeolite catalyst comprises sodalite cages and supercages, aSi/Al molar ratio of 20 or less, less than 0.5 weight percent alkalimetals; and up to 5 wt % of Pt, Pd and or Nickel, wherein, at steadystate, at least 90% of the butenes (or at least 90% of the C2 to C5olefins) are converted to products and wherein the Research OctaneNumber (RON) remains above 92; and conducting the process for a catalystage of 2.5 or greater over the same catalyst; and wherein steady statemeans that the selectivity to C8 isomers changes by 10% or less over theentire period that the catalyst age is determined.
 2. The method ofclaim 1, comprising: passing a feed mixture consisting of excessisobutane and butenes (or C2 to C5) olefins into a reaction chamber;wherein the reaction chamber comprises a crystalline zeolite catalyst;wherein the crystalline zeolite catalyst comprises sodalite cages andsupercages, a Si/Al molar ratio of 20 or less, less than 0.5 weightpercent alkali metals, and rare earth elements in the range of 10 to 35wt %; and up to 5 wt % of Pt, Pd and or Nickel, wherein, at steadystate, at least 90% of the butenes (or at least 90% of the C2 to C5olefins) are converted to products and wherein the Research OctaneNumber (RON) remains above 92; and conducting the process for a catalystage of 2.5 or greater over the same catalyst; and wherein steady statemeans that the selectivity to C8 isomers changes by 10% or less over theentire period that the catalyst age is determined.
 3. The method ofclaim 1 wherein the catalyst is regenerated in a flowing gas stream thatis essentially hydrogen at a temperature of at least 250° C.; andwherein the catalyst comprises 0.1 wt % to 5 wt % of an element selectedfrom the group consisting of Pt, Pd, Ni, and combinations thereof. 4.The method of claim 1 wherein the method is run continuously for acatalyst age of 2-3.5 without regenerating the catalyst.
 5. The methodof claim 1 wherein the reaction chamber comprises a packed catalyst bed.6. The method of claim 5 comprising a recycle stream such that thecatalyst bed I/O is greater than
 300. 7. The method of claim 1 whereinC8 selectivity is at least 50%.
 8. The method of claim 7 wherein the C2to C5 olefin consists essentially of mixed butenes.
 9. The method ofclaim 8 conducted at a temperature between 45 and 90° C.
 10. The methodof claim 1 wherein the C2 to C5 olefin contains less than 50 ppm water.11. The method of claim 1 comprising conducting the process for acatalyst age of 2.5 or a catalyst age of 3.0.
 12. The method of claim 11wherein the crystalline zeolite structure is zeolite X.
 13. A method ofalkylating isobutane, comprising: passing a feed mixture consisting ofexcess isobutane in admixture with butenes (or C2 to C5) olefins into areaction chamber; wherein the reaction chamber comprises a crystallineβ-zeolite catalyst; wherein the crystalline β-zeolite catalyst comprisesa Si/Al molar ratio of 20 or less, less than 0.5 weight percent alkalimetals; and up to 5 wt % of Pt, Pd and or Nickel, wherein, at steadystate, at least 90% of the butenes (or at least 90% of the C2 to C5olefins) are converted to products and wherein the Research OctaneNumber (RON) remains above 92; and conducting the process for a catalystage of 2.5 or greater over the same catalyst; and wherein steady statemeans that the selectivity to C8 isomers changes by 10% or less over theentire period that the catalyst age is determined.
 14. The method ofclaim 1 conducted at a pressure of 250 to 400 psig.
 15. The method ofany of claim 1 wherein the catalyst is regenerated with a flow of gaswhich is at least 50 volume % hydrogen.
 16. The method of claim 1wherein the C2 to C5 olefin contains less than 2 wt % iso-butylene. 17.The method of claim 1 wherein the C2 to C5 olefin consists essentiallyof propylene.
 18. The method of claim 1 wherein the C2 to C5 olefin feedcontains less than 2000 ppm of butadiene.
 19. The method of claim 1wherein the C2 to C5 olefin contains less than 250 ppm of mercaptans.20. The method of claim 1 wherein the C2 to C5 olefin contains less than300 ppm acetonitrile and less than 200 ppm propionitrile.